The main objectives of any method of treating liquid and solid effluents are (i) neutralizing and duly adjusting them for release and return to the environment, and (ii) obtain sub-products, recyclable, so that they may be used in the form of captive consumption and/or via commercialization to third parties.
In the exploitation of nikelferrous lateritic ores, there are several factors that influence the nickel production process by High Pressure Acid Leaching—HPAL:                geological origin of the deposit;        mineralogical composition of the ore;        particle size distribution of the ore;        operating conditions of the processing;        arrangement of the ore preparation systems, leaching, precipitation, solvent extraction and electro-refining.        
The process of high pressure acid leaching (HPAL) is recommended for predominantly limonitic ores, which have low amounts of magnesium—usually limited to 4% , maximum—because ores with high magnesium content have high consumption of sulphuric acid.
The process for production of nickel by HPAL, known in the state of the art and illustrated in FIG. 1, essentially comprises the following steps: (i) preparing the lateritic nickel ore, (ii) leaching the nickel under pressure with sulphuric acid, (iii) precipitating the nickel, (iv) re-leaching, (v) solvent extraction of the nickel and (vi) electro-refining for producing the cathode nickel (metal nickel with 99.95% purity). Due to the significant presence of cobalt in the ore, the latter will be obtained as a co-product, also in metal form.
This technology is the most appropriate process for extracting nickel and cobalt from limonitic laterites, on account of the following features:                limonitic laterites have low magnesium content and consequently, low consumption of sulphuric acid;        lower operating costs due to the low cost of sulphuric acid and its low specific consumption;        no drying ore reduction stage is needed, since the gross laterite (Run of Mine—ROM) is used in the form of slurry;        high selectivity for the metals of interest;        sulphur dioxide emissions below environmental standards;        recoveries over 90% for nickel and cobalt contained in the ore.        
Pressure leaching generally occurs in titanium-coated autoclaves, at temperatures in the range of 245 to 270° C. In this process the slurry of the autoclave feed contains approximately 40 to 45% solids, being previously heated with steam. In some ores, due to the nature of the clayey-ores present, this concentration may be limited from 25 to 30%. The level of thickening of the slurry significantly affects the capacity of the autoclave, which comprising a rather high capital cost equipment. The leaching mechanism involves acid dissolution at high temperature of the nickel and cobalt contained in the matrix of the host minerals. Under these conditions, there occurs the dissolution of the iron minerals, followed by the formation of sulfides, which, under high temperature conditions, react with water to form hematite and, consequently, regenerate the sulphuric acid:2FeOOH+3H2SO4=Fe2(SO4)3+4H2OFe2Si4O10(OH)2+3H2SO4=Fe2(SO4)3+4SiO2+4H2OFe2(SO4)3+2 H2O=2Fe(OH)SO4+H2SO42Fe(OH)SO4+H2O=Fe2O3+2H2SO4 
The extraction levels of this process reach values of 92 to 96% for nickel and 90 to 92% for the cobalt. Usually, to obtain this degree of extraction, the reaction slurry after the chemical attack should present a residual free acid concentration of 30 to 50 g/L.
After leaching, the slurry from the autoclave is depressurized and cooled in expansion chambers (“flash vessels”), to approximately 100° C., the remaining solids being separated from the liquid phase. Solid-liquid separation is performed in decanters operating in counter current (CCD), generating a liquor-loaded with sulfides of nickel, cobalt, magnesium, manganese, zinc, copper, iron and other metals. Nickel and cobalt present in the liquor are then precipitated as sulfides (using H2S), carbonates (using ammonium carbonate) or as hydroxides, using magnesia—MgO. These intermediate products usually have contents of 55% (Ni+Co) for the case of sulfides (MSP—Mixed Sulfide Precipitate) and 40 to 45% (Ni+Co) for the case of hydroxides (MHP—Mixed Hidroxide Precipitate). It is also possible to recover these metals through solvent extraction applied directly to the liquor from the decanting system. It should be noted that the MSP process makes it possible to obtain a product with a higher content of valuable metals and lowest level of contamination of manganese, magnesium and sulfides. However, production by the MSP route involves high capital cost for auxiliary installations, since there is a need for hydrogen and hydrogen sulfide units, which require sophisticated security and handlings systems for these products.
In the following step, the refining, the intermediate products (sulfides or hydroxides of nickel and cobalt) are re-leached and thus dissolved, and undergo purification treatments, such as (i) solvent extraction for separating nickel and cobalt and (ii) electrolysis (electro-refining) to achieve higher degrees of purity.
In the nickel production process by the HPAL route, from ore containing silicates and magnesium carbonate, liquid effluent generation occurs in proportions of 250 to 400 m3/t Ni produced, essentially containing magnesium and sulfate and, in small amounts, cobalt, zinc, manganese, nickel, iron, chrome, among other elements. Table 1 presents the chemical composition of the effluent to be treated and shows the significant contents of sulfate and magnesium.
TABLE 1Composition of the effluent to be treatedComponentsUnitContentNippm2.00Coppm6.00Znppm0.90Mnppm40.00Feppm10.00Crg/L3.45Mgg/L18.00S04g/L75.00NH3g/L0.50pH7.5
Different types of effluent treatment with steps aimed at recovering reagents used in leaching processes and/or the recovery of metals dispersed in effluents generated by liquid treatments of minerals are known in the state of the art.
In this sense, document GB 1.520.175 describes a process of recovering metals, such as, for example, magnesium, from aqueous sulfate solutions, through the use of lime or limestone for magnesium and sulfate precipitation. In this system, magnesium is complexed in the form of hydroxide, which precipitates jointly with the calcium sulfate. The reuse of these two elements, separately, is complicated, because they have fine particle size and certain similar physical properties, limiting the use of existing industrial processes of physical separation. Additionally, recovering the sulphur contained in the calcium sulfate requires complex calcination installations, requiring intensive use of energy. In this respect, there is one major drawback, from the point of view of cost, operational facility and simplicity of installation, when compared with the process that is the object of the present invention, in which the magnesium precipitation is carried out with the use of amines.
Another process known in the state of the art is described in document US 2009/0148366, which discloses a process for recovering metals and magnesium oxide from magnesium sulfate solutions. This process makes use of crystallization of magnesium sulfate by evaporation, requiring, based on the desired degree of hydration of the sulfate, that virtually all the water contained in the effluent be evaporated. This evaporation, if vacuum is used, may occur at temperatures in the range of 70° C. to 90° C. In a next step, to use the magnesium, the magnesium sulfate precipitate should be calcinated, so as to turn it into magnesium oxide. This operation must be performed at elevated temperatures, 700° C. to 800° C., which demands intense energy consumption. To recover the sulphur in the gas resulting from this calcination, sulphur dioxide, this must first be complexed to sulphur trioxide, using a bad of catalysts, for subsequent transformation into sulphuric acid. These operations must be carried out in complex and costly and sulphuric acid plants.
Another process known in the state of the art is described in document US 2009/0180945, which discloses a system for recovering magnesium and sulfate contained in effluents from acid leaching of lateritic ores under the form of magnesium hydroxide and magnesium oxide. This process uses ammonia as precipitation agent, the sulphur being recovered in the form of ammonium sulfate. The use of ammonia major great drawback in relation to the process that is the object of the present invention, which uses amines, since ammonia is a high toxicity gas that is hard to handle, and, once combined with sulfate, does not allow regeneration thereof, whereas the amines can be handled in liquid form at ambient temperature and, chiefly, can be regenerated for reuse in the process.